Process for the production of 1,4-butanediol, γ-butyrolactone and tetrahydrofuran

ABSTRACT

A process for the production of at least one C 4  compound selected from butane-1,4-diol, γ-butyrolactone and tetrahydrofuran comprises contacting a vaporous stream containing maleic anhydride vapour, water vapour, and carbon oxides in an absorption zone with a high boiling organic solvent thereby to form a solution of maleic anhydride in the high boiling organic solvent. Maleic anhydride in this solution is reacted under esterification conditions in an esterification zone with a C 1  to C 4  alkanol to form a solution of the corresponding di-(C 1  to C 4  alkyl) maleate in the high boiling solvent. This solution of the di-(C 1  to C 4  alkyl) maleate in the high boiling solvent is contacted with a gaseous stream containing hydrogen thereby to strip di-(C 1  to C 4  alkyl) maleate therefrom and to form a vaporous stream comprising hydrogen and di-(C 1  to C 4  alkyl) maleate. Material of this vaporous stream is contacted in a hydrogenation zone under ester hydrogenation conditions in the presence of a heterogeneous ester hydrogenation catalyst thereby to convert di-(C 1  to C 4  alkyl) maleate to at least one C 4  compound selected from butane-1,4-diol, γ-butyrolactone and tetrahydrofuran which is recovered from the hydrogenation zone in the resulting product stream. The high boiling solvent has a boiling point at atmospheric pressure that is at least about 30° C. higher than that of the di-(C 1  to C 4  alkyl) maleate.

This is a 371 application of PCT/GB 97/01285 dated May 12, 1997.

This invention relates to the production of butane-1,4-diol,γ-butyrolactone and tetrahydrofuran.

Butane-1,4-diol, together with variable amounts of γ-butyrolactone andtetrahydrofuran, can be produced by hydrogenolysis of diesters of maleicacid, fumaric acid and mixtures thereof. A major use of butane-1,4-diolis as a feedstock for the plastics industry, particularly for theproduction of polybutylene terephthalate. It is also used as anintermediate for the production of γ-butyrolactone and of the importantsolvent, tetrahydrofuran.

The maleate and fumarate diesters used as feedstock for the productionof butane-1,4-diol by such a hydrogenolysis route are convenientlyprepared from maleic anhydride, which is itself produced by vapour phaseoxidation of a hydrocarbon feedstock, such as benzene, mixed C₄ olefins,or n-butane, in the presence of a partial oxidation catalyst. In thepartial oxidation of benzene there is typically used a supportedvanadium pentoxide catalyst promoted with MoO₃ and possibly otherpromoters. The reaction temperature is from about 400° C. to about 455°C. and the reaction pressure is from about 1 bar to about 3 bar, whileabout 4 times the theoretical amount of air is used in order to stayoutside the explosive limits. The contact time is about 0.1 s. When thefeedstock is a mixed C₄ olefin feedstock, i.e. a mixed butenesfeedstock, then the partial oxidation catalyst may be vanadium pentoxidesupported on alumina. Typical reaction conditions include use of atemperature of from about 425° C. to about 485° C. and a pressure offrom about 1.70 bar to about 2.05 bar. The volume ratio of air tobutenes may be about 75:1 in order to stay below explosive limits.Alternatively it is possible, according to more modern practice, todesign the plant so that satisfactory safe operation can be achieved,despite the fact that the feed mixture of air and butenes is within theflammable limits. In the case of n-butane as feedstock, the catalyst istypically vanadium pentoxide and the reaction conditions include use ofa temperature of from about 350° C. to about 450° C. and a pressure offrom about 1 bar to about 3 bar. The air:n-butane volume ratio may beabout 20:1, even though this may be within the flammable limits. Onedesign of reactor for such partial oxidation reactions comprisesvertical tubes surrounded by a jacket through which a molten salt iscirculated in order to control the reaction temperature.

In each case a hot vaporous reaction mixture is recovered from the exitend of the reactor which comprises maleic anhydride vapour, watervapour, carbon oxides, oxygen, nitrogen, and other inert gases, besidesorganic impurities such as formic acid, acetic acid, acrylic acid, andunconverted hydrocarbon feedstock.

One way of recovering maleic anhydride from such a reaction mixture isto cool it to about 150° C. using a steam-producing stream and then tocool it further to about 60° C. by cooling it against water in order tocondense part of the maleic anhydride, typically about 30% to about 60%of the maleic anhydride present. The remainder of the stream is thenscrubbed with water.

Scrubbing with water or with an aqueous solution or slurry is described,for example, in U.S. Pat. No. 2,638,481. Such scrubbing results inproduction of a solution of maleic acid which is then dehydrated, bydistilling with xylene, for example, so as to remove the water andre-form the anhydride. A disadvantage of such a procedure, however, isthat an unacceptable proportion of the product remains in the vapourphase. In addition, some of the maleic acid is inevitably isomerised tofumaric acid. The byproduct fumaric acid represents a loss of valuablemaleic anhydride and is difficult to recover from the process systemsince it tends to form crystalline masses which give rise to processproblems.

Because of this isomerisation problem a variety of other anhydrousscrubbing liquids have been proposed. For example, dibutyl phthalate hasbeen proposed as scrubbing liquid in GB-A-727828, GB-A-763339, andGB-A-768551. Use of dibutyl phthalate containing up to 10 weight %phthalic anhydride is suggested in U.S. Pat. No. 4,118,403. U.S. Pat.No. 3,818,680 teaches use of a normally liquid intramolecular carboxylicacid anhydride, such as a branched chain C₁₂₋₁₅ -alkenyl substitutedsuccinic anhydride, for absorption of maleic anhydride from the reactionmixture exiting the partial oxidation reactor. Tricresyl phosphate hasbeen proposed for this purpose in FR-A-1125014. Dimethyl terephthalateis suggested for this duty in JP-A-32-8408 and dibutyl maleate inJP-A-35-7460. A high molecular weight wax as scrubbing solvent is taughtin U.S. Pat. No. 3,040,059, while U.S. Pat. No. 2,893,924 proposesscrubbing with diphenylpentachloride. Use of an aromatic hydrocarbonsolvent having a molecular weight between 150 and 400 and a boilingpoint above 140° C. at a temperature above the dew point of water in thevaporous reaction mixture, for example dibenzylbenzene, is suggested inFR-A-2285386. Absorption of maleic anhydride from the vaporous partialoxidation reaction mixture in dimethylbenzophenone followed bydistillation is described in U.S. Pat No. 3,850,758.Polymethylbenzophenones, at least a portion of which contain at least 3methyl groups, can be used as liquid absorbent for maleic anhydrideaccording to U.S. Pat. No. 4,071,540. Dialkyl phthalate esters having C₄to C₈ alkyl groups and a total of 10 to 14 carbon atoms in both alkylgroups are proposed for absorption of maleic anhydride from the reactionmixture in U.S. Pat. No. 3,891,680. An ester of a cycloaliphatic acid,for example dibutyl hexahydrophthalate, is suggested as absorptionsolvent for maleic anhydride in ZA-A-80/1247.

It has also been proposed to effect direct condensation of maleicanhydride from the reaction mixture exiting the partial oxidationreactor. However, this procedure is inefficient because an unacceptableproportion of the maleic anhydride remains in the vapour phase.

The maleic anhydride product recovered following condensation or byscrubbing or absorption and distillation is then reacted with a suitableC₁ to C₄ alkanol, such as methanol or ethanol, to yield thecorresponding di-(C₁ to C₄ alkyl maleate. This di-(C₁ to C₄ alkyl)maleate may contain a minor amount of the corresponding di-(C₁ to C₄alkyl) fumarate, besides traces of the corresponding mono-(C₁ to C₄alkyl) maleate and/or fumarate. It is then subjected to hydrogenolysisto yield a mixture of butane-1,4-diol, together with variable amounts ofγ-butyrolactone and tetrahydrofuran, depending upon the hydrogenolysisconditions that are selected, and of the C₁ to C₄ alkanol which can berecycled to produce further di-(C₁ to C₄ alkyl) maleate.

Processes and plant for the production of dialkyl maleates from maleicanhydride are described, for example, in U.S. Pat. No. 4,795,824 and inWO-A-90/08127. This last mentioned document describes a column reactorcontaining a plurality of esterification trays each having apredetermined liquid hold-up and containing a charge of a solidesterification catalyst, such as an ion exchange resin containingpendant sulphonic acid groups. A liquid phase containing, for example, acarboxylic acid component flows down the column from one esterificationtray to the next lower one against an upflowing stream of vapour of thelower boiling component of the esterification reagents, typically the C₁to C₄ alkanol. Water of esterification is removed from the top of thecolumn reactor in a vapour stream, while ester product is recovered fromthe sump of the reactor. As the liquid flows down the trays itencounters progressively drier reaction conditions and theesterification reaction is driven further towards 100% ester formation.This column reactor may be followed by a polishing reactor operatingunder liquid phase reaction conditions, the ester-containing stream fromthe bottom of the column reactor being admixed with further C₁ to C₄alkanol prior to admission to the polishing reactor. When used for theproduction of a di-(C₁ to C₄ alkyl) maleate, the column reactor can bepreceded by a non-catalytic monoesterification reactor in which maleicanhydride is reacted with the C₁ to C₄ alkanol in the absence of anadded catalyst to form the mono-(C₁ to C₄ alkyl) maleate.

The hydrogenation of dialkyl maleates to yield butane-1,4-diol isdiscussed further in U.S. Pat. Nos. 4,584,419, 4,751,334, andWO-A-88/00937, the disclosures of all of which are herein incorporatedby reference.

It would be desirable to simplify the production of butane-1,4,-diol,γ-butyrolactone and tetrahydrofuran, from maleic anhydride by the di-(C₁to C₄ alkyl) maleate hydrogenolysis route. In particular it would bedesirable to reduce the capital cost of construction of such a plant andalso to reduce its running costs, thereby making butane-1,4-diol,γ-butyrolactone and tetrahydrofuran more readily available.

It is accordingly an object of the present invention to simplify theproduction of butane-1,4,-diol, γ-butyrolactone and tetrahydrofuran frommaleic anhydride by the di-(C₁ to C₄ alkyl) maleate hydrogenolysisroute. A further object is to reduce the capital cost of construction ofsuch a plant by reducing significantly the numbers of distillationcolumns and the amount of other equipment required. It further seeks toreduce the running costs of a butane-1,4-diol production plant, therebymaking butane-1,4-diol, γ-butyrolactone and tetrahydrofuran more readilyavailable.

According to the present invention there is provided a process for theproduction of at least one C₄ compound selected from butane-1,4-diol,γ-butyrolactone and tetrahydrofuran, which includes the step ofhydrogenation in the vapour phase of a di-(C₁ to C₄ alkyl) maleate inthe presence of a particulate ester hydrogenation catalyst, whichprocess comprises:

(a) contacting a vaporous stream containing maleic anhydride vapour,water vapour, and carbon oxides in an absorption zone with a highboiling organic solvent having a boiling point at atmospheric pressurewhich is at least about 30° C. higher than that of the di-(C₁ to C₄alkyl) maleate thereby to form a solution of maleic anhydride in thehigh boiling organic solvent;

(b) recovering from the absorption zone a waste gas stream;

(c) reacting maleic anhydride in the solution of maleic anhydride ofstep (a) under esterification conditions in an esterification zone witha C₁ to C₄ alkanol to form the corresponding di-(C₁ to C₄ alkyl)maleate;

(d) recovering from the esterification zone a solution of the di-(C₁ toC₄ alkyl) maleate in the high boiling solvent;

(e) contacting the solution of the di-(C₁ to C₄ alkyl) maleate in thehigh boiling solvent with a gaseous stream containing hydrogen therebyto strip di-(C₁ to C₄ alkyl) maleate therefrom and to form a vaporousstream comprising hydrogen and di-(C₁ to C₄ alkyl) maleate;

(f) contacting material of the vaporous stream of step (e) in ahydrogenation zone under ester hydrogenation conditions in the presenceof a heterogeneous ester hydrogenation catalyst thereby to convertdi-(C₁ to C₄ alkyl) maleate to at least one C₄ compound selected frombutane-1,4-diol, γ-butyrolactone and tetrahydrofuran; and

(g) recovering from the hydrogenation zone a product stream containingsaid at least one C₄ compound.

Preferably in such a process the C₁ to C₄ alkanol is methanol or ethanoland the di-(C₁ to C₄ alkyl) maleate is dimethyl maleate or diethylmaleate. The use of methanol as the C₁ to C₄ alkanol and of dimethylmaleate as the di-(C₁ to C₄ alkyl) maleate is especially preferred.

The vaporous stream of step (a) of the process of the invention ispreferably produced by partial oxidation of a hydrocarbon feedstock inthe presence of a partial oxidation catalyst using molecular oxygen,typically in the form of air. The hydrocarbon feedstock can be benzene,or a mixed C₄ olefin stream, but is most preferably n-butane. The use ofn-butane as hydrocarbon feedstock is currently preferred upon thegrounds of cost since it is a cheaper feedstock than benzene or butenes.Hence in the process of the invention the feedstock used for productionof the maleic anhydride containing vaporous stream of step (a) is mostpreferably n-butane and the catalyst is preferably vanadium pentoxide.Typical partial oxidation conditions in this case include use of atemperature of from about 350° C. to about 450° C. and a pressure offrom about 1 bar to about 3 bar, an air to n-butane ratio of from about15:1 to about 50:1, e.g. about 20:1 and a partial oxidation catalystcomprising vanadium pentoxide; the contact time is typically from about0.01 s to about 0.5 s, e.g. about 0.1 s.

Partial oxidation of the hydrocarbon feedstock is conveniently conductedin a reactor which comprises vertical tubes surrounded by a jacketthrough which a molten salt is circulated in order to control thereaction temperature. The vaporous stream from the partial oxidationreactor can then be cooled by external cooling with boiler feed water toraise steam, and possibly also by further external cooling with coolingwater to a temperature in the range of from about 60° C. to about 160°C.

In step (a) of the process of the invention the vaporous maleicanhydride stream is preferably contacted with the high boiling solventat a temperature in the range of from about 60° C. to about 160° C.,preferably from about 80° C. to about 120° C., and at a pressure of fromabout 1 bar to about 3 bar so as to form a solution comprising maleicanhydride in the high boiling solvent. The contacting can be carried outby bubbling the vaporous stream through a body of the solvent.Alternatively the solvent can be sprayed into the vaporous stream.Countercurrent contacting devices can also be employed wherein theascending vaporous stream is contacted by a descending stream of solventin a gas-liquid contacting device, such as a packed scrubber tower or ascrubber tower provided with trays. In this step the solvent willtypically be at a lower temperature than the vaporous stream so that thelatter is cooled.

In the resulting solution of maleic anhydride in the high boilingsolvent the concentration of maleic anhydride in the high boilingsolvent may range from about 100 g/l to about 400 g/l.

The high boiling solvent has a boiling point at atmospheric pressurethat is at least about 30° C. higher than that of the di-(C₁ to C₄alkyl) maleate. The solvent should be selected so that it does not reactsignificantly with maleic anhydride under conditions used in thecontacting step (a) or the esterification step (c). Hence it ispreferably inert under the scrubbing conditions of step (a) as well asunder the esterification conditions used in step (c). It should also besubstantially inert under the hydrogenation conditions of step (f).

As examples of suitable high boiling solvents there can be mentioneddibutyl phthalate; tricresyl phosphate; dibutyl maleate; a highmolecular weight wax; an aromatic hydrocarbon solvent having a molecularweight between 150 and 400 and a boiling point above 140° C., such asdibenzylbenzene; and dialkyl phthalate esters having C₄ to C₈ alkylgroups and a total of 10 to 14 carbon atoms in both alkyl groups. Whenthe solvent used is an ester it is preferred that the alkyl moiety insuch an ester shall be derived from the same alkanol as the C₁ to C₄alkanol used in the esterification step (c). In this way anytransesterification reactions that may occur do not give rise toadditional esters. Thus when the alkanol used is methanol and thedialkyl maleate is dimethyl maleate, any ester used as the high boilingsolvent is preferably also a methyl ester. Examples of such methylesters which can be used as the high boiling solvent include dimethylphthalate, dimethyl esters of other aromatic acids, such as dimethyl2,3-naphthalenedicarboxylate, diesters of cyclic aliphatic diacids, suchas dimethyl 1,4-cyclohexanedicarboxylate, and methyl esters of longchain fatty acids containing, for example, from 14 to 30 carbon atoms.Other solvents that can be used include high boiling ethers such asdimethyl ethers of polyethylene glycols of appropriate molecular weight,such as tetraethyleneglycol dimethyl ether.

Another desirable quality of the high boiling solvent is that it shouldbe essentially water insoluble and/or essentially incapable ofdissolving water.

The high boiling solvent used in step (a) conveniently comprisesmaterial resulting from the hydrogen stripping step (e). Hence it maycontain already some di-(C₁ to C₄ alkyl) maleate.

Provided that appropriate conditions are adopted in step (a), the gasstream recovered in step (b) of the process of the invention can beessentially free from maleic anhydride.

Esterification of the maleic anhydride with the C₁ to C₄ alkanol iseffected in step (c) in an esterification zone. This may comprise anon-catalytic reactor in which the maleic anhydride in the solution inthe high boiling solvent undergoes reaction in the absence of addedcatalyst with the C₁ to C₄ alkanol to form the corresponding mono-(C₁ toC₄ alkyl) maleate. The reaction is:- ##STR1## where R is a C₁ to C₄alkyl radical. Some conversion of the mono-(C₁ to C₄ alkyl) maleate tothe corresponding di-(C₁ to C₄ alkyl) maleate may also occur. Thereaction concerned is:- ##STR2## where R is as defined above.

Such a non-catalytic reactor can be operated under monoesterificationconditions which typically comprise use of a temperature of from about65° C. to about 260° C. and a pressure of from about 1 bar to about 50bar. This can be followed by a catalytic esterification stage. Forexample, the catalytic esterification stage may comprise a plurality ofstirred tank reactors such as is disclosed in U.S. Pat. No. 4,795,824.Preferably, however, the catalytic esterification stage comprises acolumn reactor of the type disclosed in WO-A-90/08127. In this case thenon-catalytic monoesterification stage may comprise a stirred tankreactor or a column reactor containing one or more trays which do notcontain any esterification catalyst and which is fed from the bottomwith methanol or other C₁ to C₄ alkanol vapour, while the maleicanhydride solution from step (a) is fed downward through the columnreactor.

If the catalytic esterification stage comprises a column reactor of thetype disclosed in WO-A-90/08127, then the solution of maleic anhydride(or a solution comprising the corresponding mono-(C₁ to C₄ alkyl)maleate, if a separate monoesterification stage is used) in the highboiling liquid is fed to the top esterification tray of the columnreactor, while an excess of C₁ to C₄ alkanol vapour is fed to the bottomof the reactor.

In the column reactor the esterification trays each hold a charge of asolid esterification catalyst. Each tray has a vapour upcomer means topermit vapour to enter the tray from below and to agitate the mixture ofliquid and solid esterification catalyst in a zone of turbulence on thetray and to keep the catalyst particles in suspension. In order avoidthe danger of "hot spots" forming on the tray through formation ofpockets of settled catalyst particles, the floor of each tray ispreferably designed so as to slope towards the zone of turbulence at aslope which exceeds the angle of repose of the catalyst particles underthe liquid. In addition each esterification tray has a downcomer meanswhich permits liquid, but not catalyst particles, to flow down from thattray to the next lower one. Such a downcomer means will usually beprovided with a screen to prevent catalyst particles passing downwardlytherethrough.

Typical reaction conditions in the column reactor include use of atemperature and pressure under which the C₁ to C₄ alkanol distils. Suchtemperature and pressure conditions will vary in dependence upon the C₁to C₄ alkanol selected but will typically include use of a temperatureof from about 65° C. to about 135° C. and a pressure of from about 1 barto about 3 bar. A typical solid esterification catalyst is the ionexchange resin sold under the designation Amberlyst™ 16 by Rohm and Haas(U.K.) Limited of Lennig House, 2 Mason's Avenue, Croydon CR9 3NB.

In passing up the column from one esterification tray to the next higherone, the upflowing C₁ to C₄ alkanol vapour carries with it water ofesterification. Thus the di-(C₁ to C₄ alkyl) maleate-containing liquidpassing down the column reactor from one esterification tray to the nextlower one encounters drier and drier conditions as it proceeds down thecolumn. In this way the esterification reaction leading to formation ofthe di-(C₁ to C₄ alkyl) maleate is driven further and further towards100% conversion to the di-(C₁ to C₄ alkyl) maleate.

Any byproduct acid, such as acetic acid or acrylic acid, that is alsopresent in the vaporous stream from the partial oxidation reactor,together with any maleic acid or fumaric acid present in the solutionsupplied to the esterification zone, will undergo conversion to thecorresponding C₁ to C₄ alkyl ester or diester, as the case may be.

The vapour phase stream emerging from the topmost esterification traycomprises C₁ to C₄ alkanol vapour and water vapour; it may furtherinclude traces of minor byproducts such as the di-(C₁ to C₄ alkyl)ether, besides traces of the di-(C₁ to C₄ alkyl) maleate and of the C₁to C₄ alkyl acrylate. A further additional tray or trays may be providedabove the uppermost esterification tray to act as a form of washingcolumn in order to return di-(C₁ to C₄ alkyl) maleate to theesterification trays. The resulting vapour stream, which is nowessentially free from di-(C₁ to C₄ alkyl) maleate, exits the top of thecolumn.

From the bottom of the column reactor there is recovered a liquid streamcomprising a solution of the di-(C₁ to C₄ alkyl) maleate in the highboiling solvent. This is essentially acid free. If desired this liquidcan be admixed with additional C₁ to C₄ alkanol and passed through apolishing reactor containing a bed of solid esterification catalystoperating under liquid phase operating conditions. Such conditionstypically include use of a temperature of from about 65° C. to about135° C. and a pressure of from about 1 bar to about 3 bar. A typicalsolid esterification catalyst is the ion exchange resin sold under thedesignation Amberlyst™ 16 by Rohm and Haas (U.K.) Limited of LennigHouse, 2 Mason's Avenue, Croydon CR9 3NB.

In step (e) of the process of the invention, a gas stream comprisinghydrogen is passed through the solution of the di-(C₁ to C₄ alkyl)maleate.

The hydrogen stripping step is preferably conducted substantially at orat a pressure slightly higher than the inlet pressure to the esterhydrogenation zone. The hydrogen stripping step is similarly preferablyconducted at substantially the desired inlet temperature to thehydrogenation step or a little below this temperature, for example fromabout 5° C. to about 20° C. below this temperature. Then the temperaturecan be raised to the desired inlet temperature by admixture of furtherhot hydrogen-containing gas which has the additional benefit of dilutingthe vaporous ester-containing stream and thereby ensuring that it is ata temperature above its dew point, preferably at least about 5° C.higher than its dew point.

The hydrogenation step is advantageously conducted in the vapour phase,using a heterogeneous ester hydrogenation catalyst. Typical esterhydrogenation catalysts include reduced promoted copper catalysts, forexample reduced copper chromite catalysts such as that sold under thedesignation PG 85/1 by Davy Process Technology Limited of 30 EastbourneTerrace, London W2 6LE.

The catalyst particles preferably have a particle size in the range offrom about 0.5 mm to about 5 mm. The particles may be of any convenientshape, e.g. spheres, pellets, rings or saddles. When using a fixed bedof catalyst the reactor can be a shell-and-tube reactor, which can beoperated substantially isothermally; however, it is preferably anadiabatic reactor. The use of an adiabatic reactor is advantageous sinceits capital cost is much lower than that of a shell-and-tube reactor andit is generally much easier to charge the reactor with the chosencatalyst.

Hydrogenation is conducted at an elevated temperature of, for example,from about 150° C. to about 240° C. and at a pressure of from about 5bar to about 100 bar, preferably from about 50 bar to about 70 bar.

From the hydrogenation zone there is recovered a hydrogenation productmixture which contains, in addition to the C₁ to C₄ alkanol, alsobutane-1,4-diol, and some tetrahydrofuran and γ-butyrolactone. Even ifthe primary product of interest is butane-1,4-diol, the presence ofthese minor amounts of tetrahydrofuran and γ-butyrolactone is notdisadvantageous since these compounds are important chemicals ofcommerce and it is accordingly economic to recover them in pure form. Ifdesired, γ-butyrolactone can be recycled to the hydrogenation zone toproduce additional butane-1,4-diol. In addition the hydrogenolysisproduct mixture will normally contain minor amounts of the correspondingdi-(C₁ to C₄ alkyl) succinate, n-butanol, the corresponding dialkylalkoxysuccinate, e.g. dimethyl methoxysuccinate if the C₁ to C₄ alkanolis methanol, and water.

For further details regarding hydrogenation of a di-(C₁ to C₄ alkyl)maleate and subsequent purification of the resultant crude hydrogenationproduct mixture, reference may be made to U.S. Pat. No. 4,584,419,WO-A-86/03189, WO-A-88/00937, U.S. Pat. Nos. 4,767,869, 4,945,173,4,919,765, 5,254,758, 5,310,954, and WO-A-91/01960.

In order that the invention may be clearly understood and readilycarried into effect a plant for the production of butane-1,4-diol, aswell as some γ-butyrolactone and tetrahydrofuran, using a preferredprocess in accordance with the present invention will now be described,by way of example only, with reference to the accompanying drawing whichis a flow diagram of the plant.

Referring to the drawing, n-butane is supplied in line 1 at a pressureof from 1 to 3 bar and at a temperature of 400° C. to a partialoxidation plant 2 which is also supplied with air in line 3. Partialoxidation plant 2 is of conventional design and includes a partialoxidation reactor comprising tubes packed with a partial oxidationcatalyst consisting of vanadium pentoxide packed into tubes providedwith a jacket through which molten salt can be circulated for thepurpose of temperature control. The partial oxidation reactor isoperated at an air:n-butane feed ratio of 20:1.

A hot vaporous partial oxidation product stream is cooled by externalcooling against boiler feed water to raise steam and then againstcooling water to reduce its temperature to 138° C. It is recovered fromplant 2 in line 4. This contains 2.9% w/w maleic anhydride, 5.8% w/wwater, 1.3% w/w carbon dioxide, 1.0% w/w carbon monoxide, 0.01% w/wacetic acid, 0.0% w/w acrylic acid, 15.7% w/w oxygen, and the balanceessentially comprising nitrogen and other inert gases. It is fed to thebottom of a scrubbing tower 5, up which it passes against a downflowingspray of dimethyl phthalate which is supplied at a temperature of about68° C. from line 6. The scrubbed waste gas stream which contains 0.03%w/w maleic anhydride exits the top of scrubbing tower 5 in vent gas line7 and is passed to a waste gas burner.

From the bottom of scrubbing tower 5 there is recovered a liquid streamin line 8 which comprises a solution of approximately 22% w/w maleicanhydride and 0.04% w/w acrylic acid in dimethyl phthalate. This issupplied to the top of a column reactor of the type described inWO-A-90/08127. This comprises a number of esterification trays mountedone above the other, each containing a charge of a solid esterificationcatalyst, such as Amberlyst™ 16 resin, and each having a vapour upcomerfor upflowing vapour and a liquid downcomer to permit liquid to flowdown the column from one esterification tray to the next lower one.Methanol vapour is supplied to the bottom of column reactor by way ofline 10. Water of esterification is removed in the vapour stream exitingthe column reactor in line 11. Column reactor 9 is operated at atemperature of from about 110° C. to about 125° C. and at a pressure offrom about 1 bar to about 3 bar. The residence time in the columnreactor is about 3 hours. Normally the temperature on the top tray willbe somewhat higher (e.g. about 125° C.) than that on the lowermost tray(e.g. about 115° C.).

A solution containing about 250 g/l dimethyl maleate in dimethylphthalate is withdrawn from the bottom of column reactor 9 in line 12and pumped to near the top of a stripping column 13 which is operated ata temperature of 170° C. and a pressure of 885 psia (61.02 bar). Column13 has a number of distillation trays above the point of injection ofthe dimethyl maleate solution into column 13 so as to reduce carryoverof the high boiling solvent dimethyl phthalate in the overhead streamfrom column 13. The solution of dimethyl maleate in dimethyl phthalateflows down stripping column 13 against an upflowing stream of hydrogenfrom line 14. The stripped dimethyl phthalate is recycled from thebottom of stripping column 13 by way of line 6 to the top of scrubbingtower 5. From the top of stripping column 13 there emerges in line 15 anear saturated vapour mixture stream comprising dimethyl maleate inhydrogen, with a hydrogen:dimethyl maleate molar ratio of about 320:1.This vapour mixture stream is at a temperature of from about 180° C. toabout 195° C. and at a pressure of 62 bar. It is diluted with furtherhot hydrogen at a temperature of from about 180° C. to about 195° C. toyield a vaporous stream with a hydrogen:dimethyl maleate molar ratio ofabout 350:1 and is at least about 5° C. above its dew point.

This vaporous mixture passes onwards in line 17 to hydrogenation plant18 which includes an adiabatic reactor packed with a reduced copperchromite catalyst and operated at an inlet temperature of 173° C., aninlet pressure of 885 psia (61.02 bar), and an exit temperature of 190°C. The dimethyl maleate feed rate corresponds to a liquid hourly spacevelocity of 0.5 h⁻¹. The plant also includes a purification section inwhich the crude hydrogenation product mixture is distilled in severalstages to yield pure butane-1,4-diol in line 19. Lines for separaterecovery of γ-butyrolactone and tetrahydrofuran are indicated at 20 and21 respectively.

We claim:
 1. A process for the production of at least one C₄ compoundselected from butane-1,4-diol, γ-butyrolactone and tetrahydrofuran,which includes the step of hydrogenation in the vapour phase of a di-(C₁to C₄ alkyl) maleate in the presence of a particulate esterhydrogenation catalyst, which process comprises:(a) contacting avaporous stream containing maleic anhydride vapour, water vapour, andcarbon oxides in an absorption zone with a high boiling organic solventhaving a boiling point at atmospheric pressure which is at least about30° C. higher than that of the di-(C₁ to C₄ alkyl) maleate thereby toform a solution of maleic anhydride in the high boiling organic solvent;(b) recovering from the absorption zone a waste gas stream; (c) reactingmaleic anhydride in the solution of maleic anhydride of step (a) underesterification conditions in an esterification zone with a C₁ to C₄alkanol to form the corresponding di-(C₁ to C₄ alkyl) maleate; (d)recovering from the esterification zone a solution of the di-(C₁ to C₄alkyl) maleate in the high boiling solvent; (e) contacting the solutionof the di-(C₁ to C₄ alkyl) maleate in the high boiling solvent with agaseous stream containing hydrogen thereby to strip di-(C₁ to C₄ alkyl)maleate therefrom and to form a vaporous stream comprising hydrogen anddi-(C₁ to C₄ alkyl) maleate; (f) contacting material of the vaporousstream of step (e) in a hydrogenation zone under ester hydrogenationconditions in the presence of a heterogeneous ester hydrogenationcatalyst thereby to convert di-(C₁ to C₄ alkyl) maleate to at least oneC₄ compound selected from butane-1,4-diol, γ-butyrolactone andtetrahydrofuran; and (g) recovering from the hydrogenation zone aproduct stream containing said at least one C₄ compound.
 2. A processaccording to claim 1, in which the C₁ to C₄ alkanol is methanol and thedi-(C₁ to C₄ alkyl) maleate is dimethyl maleate.
 3. A process accordingto claim 1, in which the vaporous stream of step (a) is produced bypartial oxidation of a hydrocarbon feedstock in the presence of apartial oxidation catalyst using molecular oxygen.
 4. A processaccording to claim 3, in which the hydrocarbon feedstock is n-butane. 5.A process according to claim 4, in which the partial oxidation catalystcomprises vanadium pentoxide and in which the partial oxidationconditions include use of a temperature of from about 350° C. to about450° C., a pressure of from about 1 bar to about 3 bar, an air ton-butane ratio of from about 15:1 to about 50:1 and a contact time offrom about 0.01 s to about 0.5 s.
 6. A process according to claim 1, inwhich in step (a) the vaporous maleic anhydride stream is contacted withthe high boiling solvent at a temperature in the range of from about 60°C. to about 160° C. and at a pressure of from about 1 bar to about 3 barso as to form a solution comprising maleic anhydride in the high boilingsolvent.
 7. A process according to claim 6, in which the contacting stepis carried out in a countercurrent contacting device wherein theascending vaporous stream is contacted by a descending stream of solventin a gas-liquid contacting device.
 8. A process according to claim 1, inwhich the solvent is an alkyl ester whose alkyl moiety is derived fromthe same alkanol as the C₁ to C₄ alkanol used in the esterification step(c).
 9. A process according to claim 1, in which the C₁ to C₄ alkanol ismethanol, the di-(C₁ to C₄ alkyl) maleate is dimethyl maleate, and thehigh boiling solvent is also a methyl ester.
 10. A process according toclaim 9, in which the methyl ester is dimethyl phthalate.
 11. A processaccording to claim 9, in which the methyl ester is a methyl ester ormixture of methyl esters of a long chain fatty acid or acids containingfrom 14 to 30 carbon atoms.
 12. A process according to claim 1, in whichthe high boiling solvent is a dimethyl ether of a polyethylene glycol.13. A process according to claim 1, in which the high boiling solventused in step (a) comprises recycled material resulting from the hydrogenstripping step (e).
 14. A process according to claim 1, in which theesterification zone comprises a non-catalytic reactor in which themaleic anhydride in the solution in the high boiling solvent undergoesreaction in the absence of added catalyst with the C₁ to C₄ alkanol toform the corresponding mono-(C₁ to C₄ alkyl) maleate.
 15. A processaccording to claim 1, in which the catalytic esterification stagecomprises a column reactor provided with a plurality of esterificationtrays each of which holds a charge of a solid esterification catalyst,has a vapour upcomer means to permit vapour to enter the tray from belowand to agitate the mixture of liquid and solid esterification catalystin a zone of turbulence on the tray and to keep the catalyst particlesin suspension, and a downcomer means which permits liquid, but notcatalyst particles, to flow down from that tray to the next lower one,the column reactor being supplied beneath the lowermost esterificationtray with a stream of C₁ to C₄ alkanol vapour and to an upperesterification tray with a solution in the high boiling solventcomprising a material selected from maleic anhydride, a mono-(C₁ to C₄alkyl) maleate wherein the C₁ to C₄ alkyl group is derived from the C₁to C₄ alkanol, and a mixture thereof.
 16. A process according to claim15, in which the floor of each tray slopes towards the zone ofturbulence at a slope which exceeds the angle of repose of the catalystparticles under the liquid.
 17. A process according to claim 1, in whichthe esterification zone comprises an autocatalytic esterification zonewherein the esterification conditions include use of a temperature offrom about 70° C. to about 250° C., a pressure of from about 1 bar toabout 50 bar and wherein maleic anhydride is converted by reaction withC₁ to C₄ alkanol at least in part to the corresponding mono-(C₁ to C₄alkyl) maleate.
 18. A process according to claim 1, wherein theesterification zone includes a catalytic esterification zone wherein theesterification conditions include use of a temperature of from about 65°C. to about 135° C. and of a solid esterification catalyst comprising anion exchange resin containing pendant sulphonic acid groups.
 19. Aprocess according to claim 1, in which the hydrogen stripping step isconducted at substantially the inlet pressure to the ester hydrogenationzone.
 20. A process according to claim 1, in which the hydrogenstripping step is conducted at a temperature in the range of from theinlet temperature to the hydrogenation zone to about 20° C. below theinlet temperature to the hydrogenation zone.
 21. A process according toclaim 1, in which the hydrogenation step is conducted in the vapourphase using a reduced promoted copper catalyst at a temperature of fromabout 150° C. to about 240° C. and at a pressure of from about 5 bar toabout 100 bar.
 22. A process according to claim 1, in which there isrecovered from the hydrogenation zone a hydrogenation product mixturewhich contains, in addition to butane-1,4-diol and the C₁ to C₄ alkanol,also minor amounts of tetrahydrofuran and γ-butyrolactone.
 23. A processaccording to claim 22, in which the hydrogenation product mixture ispurified by distillation in one or more stages, including distillationin a "light ends" column to separate overhead the volatile components ofthe mixture including tetrahydrofuran, the C₁ to C₄ alkanol, water, andn-butanol.
 24. A process according to claim 23, in which the bottomsproduct from the "light ends" column is further purified by distillationin one or more stages to yield pure butane-1,4-diol.